02 BOPs / Woods D.R 2008 rules-of-thumb-in-Engineering-practice (epdf.tips)

6.9 PFTR: Fixed Bed Catalyst in Tube or Vessel: Adiabatic 231

Concentration gradient within the catalyst and in the external bulk liquid phase:

3.Internal concentration gradient depends on the reaction.

4.External concentration gradient may be high because of slow mass transfer or fast reaction.

For IX resin, height/diameter = 0.5. If the main reactant undergoes 90 % conversion with the bed height/catalyst particle diameter = 10 000, then the reaction is not mass transfer controlled. Since this height/diameter ratio is usually not used, check if mass transfer controls.

Related topic liquid adsorption, Section 4.12 and IX, Section 4.13. Gas–liquid downflow over fixed catalyst, see Trickle reactors, Section 6.17.

x Good Practice

The target inlet gas temperature should be such that the initial rate of reaction is in the following target ranges:

If the reaction is exothermic then set the inlet temperature such that the target rate of reaction is 0.2 q 10–5–0.5 q 10–5 mol/s.g catalyst. For a catalyst with bulk density 1.4 Mg/m3 this is about 1 1/s.

If the reaction is endothermic then set the inlet temperature such that the target rate of reaction is 0.4 q 10–4 –8 q 10–4 .

Increase the reaction temperature gradually to offset catalyst decay.

x Trouble Shooting1)

Gas-catalytic reactions. Temperature and pressure drop across bed are usually key variables. When a hot spot develops, it usually develops at the front end of the bed and gradually moves through the bed. It may take three to four weeks to travel through the full bed. If the hot spot is 100–200 hC above normal, then usually carbon is deposited and the catalyst is irrevocably damaged. Temperature control is critical for exothermic reactions. “Dp rapidly increases”: emergency shutdown?. “Pressure surge”: possible shutdown?/[runaway reactor]*.

“Rapid decline in conversion”: unfavorable shift in equilibrium at operating temperature, for exothermic reactions/[sintering]*/[agglomeration]*/poison in new feed. “Gradual decline in conversion”: sample error/analysis error/temperature sensor error/[catalyst activity lost]*/[maldistribution]*/[unacceptable temperature profiles]*/wrong locations of feed, discharge or recycle lines/faulty design of feed and discharge ports/wrong internal baffles and internals/faulty bed voidage profiles. “Gradual decline in conversion and axial temperature constant with depth of region increasing with time”: [poisoned catalyst]*. “Gradual decline in conversion and axial temperatures I usual”: [poisoned catalyst]*. “Gas exit concentration of reactants high”: sample error/analysis error/catalyst selectivity low/[catalyst activity lost]*.

“Exit concentration of product higher than design”: reactor leaking. “Change in prod-

1) Based on R.B. Anderson, personal commu-	and R. Gauly, Hydrocarbon Process., 1999,
nication; H.F. Rase “Fixed bed reactor design	Sept., 43–50.
and diagnostics”, 1990, Wiley and Dutta, S.

232 6 Reactors

uct distribution”: [maldistribution]*/[poisoned catalyst]*/feed contaminants/ change in feed/change in temperature settings.

“Temperature runaways”: [temperature hot spots]*/[reactor instability]*. “Pressure and bed temperature and reactor unsteady”: water in feed/[maldistribution]*. “Local high temperature/hot spot with T i 100 hC above normal”: [maldistribution of gas flow]*/instrument error/extraneous feed component that reacts exothermically. “Local low temperature within the bed”: [maldistribution of gas flow]*/instrument error/extraneous feed component that reacts endothermically. “Exit gas temperature too high”: instrument error/control system malfunction. “Temperature varies axially across bed”: [maldistribution]*.

“Dp higher than design”: catalyst degradation/instrument error/high gas flow/sudden coking/crud left in from construction or revamp.“Dp increasing gradually yet flowrate constant”: [coke formation]*/[dust or corrosive products from upstream processes]*.

“Startup after catalyst regeneration, conversions I standard”: [regeneration faulty]*. “Startup after catalyst replacement, poor selectivity”: bad batch of catalyst/preconditioning of catalyst faulty/temperature and pressures incorrectly set/instrument error for pressure or temperature.

“Startup after catalyst replacement, Dp I expected and conversion I standard”: [maldistribution]* and axial variation in temperature/larger size catalyst. “Startup after catalyst replacement, conversion I standard and Dp increasing”: [maldistribution and axial temperatures different]*/feed precursors present for polymerization or coking. “Startup after catalyst replacement, Dp for this batch of catalyst i previous batch”: catalyst fines produced during loading/poor loading. “Startup after catalyst replacement, conversion I specifications per unit mass of catalyst and more side reactions”:

[maldistribution]*/faulty inlet distributor/faulty exit distributor.

[Active species volatized]*: [regeneration faulty]*/faulty catalyst design for typical reaction temperature/[temperature hot spots]*.

[Agglomeration of packing or catalyst particles]*: [temperature hot spots]*.

[Attrition of the catalyst]*: flowrates i expected/catalyst too fragile. [Carbon buildup]*: [inadequate regeneration]*/[excessive carbon formed]*.

[Catalyst selectivity changes]*: [poisoned catalyst]*/feed contaminants/change in feed/change in temperature settings.

[Catalyst activity lost]*: [carbon buildup]*/[regeneration faulty]*/[sintered catalyst]*/excessive regeneration temperature/[poisoned catalyst]*/[loss of surface area]*/[agglomeration]*/[active species volatized]*.

[Excessive carbon formed]*: operating intensity above usual/feed changes/[temperature hot spots]*.

[Dust or corrosive products from upstream processes]*: in-line filters not working or not installed/dust in the atmosphere brought in with air/air filters not working or not installed.

[Loss of surface area]*: [sintered catalyst]*/[carbon buildup]*/[agglomeration]*. [Maldistribution]*: faulty flow distributor design/plugging of flow distributors with fine solids, sticky byproducts or trace polymers/[sintered catalyst particles]*/

6.10 PFTR: Multi-bed Adiabatic with Inter-bed Quench or Heating 233

[agglomeration of packing or catalyst particles]*/fluid feed velocity too high/faulty loading of catalyst bed/incorrect flow collector at outlet.

[Poisoned catalyst]*: poisons in feed/flowrate of “counterpoison” insufficient/ poison formed from unwanted reactions.

[Poisons in feed]*: depends on reaction/contamination in feed/upstream process or equipment upsets/changes in feed. Poisons for platforming include high sulfur in feed and high feed end point with upstream equipment failure being compressor failure/water upset/chloride upset.

[Reactor instability]*: control fault/poor controller tuning/wrong type of control/ feed temperature exceeds threshold.

[Regeneration doesn’t remove all carbon from the catalyst]*: regeneration temperature not hot enough/regeneration time not long enough/[maldistribution]*. [Regeneration faulty]*: temperatures too high/oxygen concentration I standard/ oxygen concentration i standard causing too rapid a burn/incorrect temperature and time so that coke left on catalyst. [regeneration doesn’t remove all carbon from the catalyst]*/excessive temperature during regeneration.

[Runaway reactor]*: feed temperature too high/[temperature hot spot]*.

[Sintered catalyst]*: temperature sensor error/[temperature hot spots]*/[maldistribution]*/temperature in reactor too high/regeneration temperature too high. [Temperature hot spots]*: bed too deep/[maldistribution]*/flowrate I design/instrument error/extraneous feed component that reacts exothermically.

6.10

PFTR: Multi-bed Adiabatic with Inter-bed Quench or Heating

x Area of Application

Phases: Gas plus solid catalyst. For fast reactions, that are strongly exothermic or endothermic. Use if the order of the reaction is positive and i 95 % conversion is the target, and for consecutive reactions with an intermediate as the target product. Used primarily for equilibrium reactions.

x Guidelines

Limit the height of the bed to keep temperature increase I 50 hC to minimize effects of radial temperature gradients (see Section 6.9). The bed can be shallow and wide. Quench can include injection of cold reactants, internal or external heat exchangers.

x Good Practice

Gas flow down through the bed. Provide good gas distribution. Increase reaction temperature gradually to offset catalyst decay.

2346 Reactors

6.11

PFTR: Fixed Bed with Radial Flow

For fast reactions, strongly exothermic or endothermic reactions.

Phases: Gas with solid catalyst. Use to minimize pressure drop limitations. Multi-staging is possible. Care is needed in sizing the gas distribution and collection. Otherwise size, cost and trouble shoot similar to Sections 6.9 and 6.10.

6.12

PFTR: Multitube Fixed Bed Catalyst or Bed of Solid Inerts: Nonadiabatic

x Area of Application

Phase: Gas–liquid plus solid catalyst

Use if the order of the reaction is positive and i 95 % conversion is the target, and for consecutive reactions with an intermediate as the target product. Exchange heat generated if the product of the Arrhenius number and the dimensionless adiabatic temperature rise i 10.

Homogeneous gas reactions requiring good temperature control via inert solids or heterogeneous gas reactions. Can provide continuous temperature control. Can be used for exothermic reactions.

x Guidelines

Shell and tube exchanger with reactants and catalyst inside the tube, 250– 400 m2/m3. Tube diameter I 50 mm. The smaller the diameter of the tubes, the larger the surface area from the tubes.

Gas with fixed bed of catalyst: Use high mass gas velocity to improve heat transfer kg/s m2 i 1.35. To ensure good gas distribution and negligible backmixing, Pe i 2; Height/catalyst particle diameter H/Dp i 100 and Dp/DI 0.10. Gas velocity 3–10 m/s; residence time 0.6–2 s. For fast reactions, catalyst pore diffusion mass transfer may control if catalyst diameter i 1.5 mm. Heat transfer coefficient: Gas at 0.1 kPa g vs. liquid: U = 0.05 kW/m2 K; gas at 20 MPa vs. liquid: U = 0.5 kW/m2 K.

Individual coefficient on shell side: coolants: boiling water, h = 1–3 kW/m2 K; boiling organic, h = 0.2–1.5 kW/m2 K; molten salts, h = 0.5–1.5 kW/m2 K. Heating agents: steam, h = 2–5 kW/m2 K; combustion gas, h = 0.01–0.03 kW/m2 K.

Liquids with fixed bed of catalyst: To minimize backmixing, Pe i 1; use H/Dp i 200 and Dp/DI 0.10. Liquid velocity 1–2 m/s; residence time 2–6 s.

Heat transfer coefficient U = 0.5 kW/m2 K. Cooling liquid vs. liquids: U = 0.2–1.2 kW/m2 K. Heating liquids via steam, U = 0.35–1.2 kW/m2 K.

x Good Practice

Gas: The target inlet gas temperature should be such that the initial rate of reaction is in the following target ranges:

6.12 PFTR: Multitube Fixed Bed Catalyst or Bed of Solid Inerts: Nonadiabatic 235

If the reaction is exothermic then set the inlet temperature such that the target rate of reaction is 1 q 10–5–2 q 10–5 mol/s g catalyst, and if endothermic then to

2q 10–4 –4 q 10–4 mol/s g catalyst.

Increase reaction temperature gradually to offset catalyst decay.

x Trouble Shooting

See Section 6.9 plus the additional considerations because of the tubes.“Dp increases dramatically, top of tubes hot, less conversion than expected”: possible shutdown?/contamination in feed/[poisoned catalyst]*.

“Gradual decline in conversion”: sample error/analysis error/temperature sensor error/[catalyst activity lost]*/[maldistribution]*/[unacceptable temperature profiles]*/[inadequate heat transfer]*/wrong locations of feed, discharge or recycle lines/faulty design of feed and discharge ports/wrong internal baffles and internals/faulty bed voidage profiles. “Exit gas temperature too high”: instrument error/control system malfunction/fouled reactor coolant tubes. “Soon after startup, temperature of tubewall near top i usual and increasing and perhaps Dp increase and less conversion than expected or operating temperatures i usual to obtain expected conversion”: inadequate catalyst regeneration/contamination in feed; for steam reforming sulfur concentration i specifications/wrong feed composition; for steam reforming: steam/CH4 I 7 to 10. “Soon after startup, temperatures over full length of some tubes i usual and perhaps Dp i or I usual and may increase with time”: faulty loading of the catalyst/[maldistribution]*. “Hot bands or stripes; perhaps Dp increase”: low ratio of steam to methane/[carbon formation; whisker type]*/wrong feed composition: for steam reforming steam/methane I 7 to 10:1. “Hot bands or stripes near top and perhaps over all tube and rapidly increasing

Dp and conversion I specifications”: [deactivated catalyst by pyrolytic coke formation]*/feed concentration wrong: for steam reforming high concentration of heavier hydrocarbons/steam to hydrocarbon ratio low/[catalyst poisoned]* by sulfur. “Temperature at inlet high and high Dp”: [ for steam reforming: steam contaminated with inorganic solids]*. “Hot bands in top 1/3 of tubes and methane i usual in exit gas and perhaps Dp increase”: contamination in feed/[poisoned catalyst]*.

“Startup after catalyst replacement, poor selectivity”: bad batch of catalyst/preconditioning of catalyst faulty/[tube walls not passified]*/temperature and pressures incorrectly set/instrument error for pressure or temperature. “Startup after catalyst replacement, increased side reactions and conversion I specification”: catalyst loading not the same in all tubes.

[Reactor instability]*: control fault/poor controller tuning/wrong type of control/insufficient heat transfer area/feed temperature exceeds threshold/coolant temperature exceeds threshold/coolant flowrate I threshold/tube diameter too large.

[Runaway reactor]*: feed temperature too high/[temperature hot spot]*/cooling water too hot/feed temperature too high.

[Tube walls not passified]*: walls activated unwanted side reactions and faulty passivation treatment/wrong passivation treatment/no passivation treatment.

2366 Reactors

6.13

PFTR: Bubble Reactor

Many different configurations fall under this general title. A bubble column is typically a tall, narrow column of liquid into which gas is sparged at the bottom. The bubbles rise through the liquid and react. The bubble column may be operated with liquid and gas flowing cocurrently or countercurrently. Packing can fill the column to create a packed bubble column. An air lift loop is a bubble column with a central draft tube. Compressed gas is injected at the bottom of the central draft tube. (Similar to a Pachuca tube reactor.) A jet loop column is a bubble column with a central draft tube with compressed gas and liquid injected at the bottom of the draft tube. The liquid is withdrawn from the annulus and pumped into the ejector below the draft tube. The power required (about 5 kW/m3) is greater than that required for either an air lift or bubble column (about 3 kW/m3). Other variations involve how the liquid recycles to the bottom. A central baffle bubble column provides a vertical axial baffle instead of a central draft tube with compressed gas being injected at one side of the baffle; liquid circulates back around the other side. An external loop bubble column is a vertical column, compressed air is injected at the bottom; an external pipe connects the liquid from the top to the bottom and this allows natural circulation. A deep shaft; tower loop with downflow air lift is a vertical column with a central draft tube. Compressed air is injected into the annulus; liquid is sparged downwards in the central annulus.

A bubble reactor is also used for aerobic reactors for the treatment of waste water. The configuration may be a basin, pond or lagoon. See Sections 6.31 and 6.32 for activated sludge reactors.

A specialized unit is an ozone generator and contact reactor.

Many other options are available for GL contacting and OTR, Sections 1.6.1 and 1.6.3 provide data to guide in selecting the options.

x Area of Application

Phases: GL, GLcS and LL. Relatively slow reactions. Use if the order of the reaction is positive and i 95 % conversion is the target, and for consecutive reactions with an intermediate as the target product. The bubble columns tend to operate isothermally and, unless heat is removed in the external loop, this configuration is not used for highly exothermic reactions.

Gas–liquid: For large liquid holdup, slow reactions that are kinetically controlled reactions that require long residence times and low viscosity liquids. Preferred if large gas volumes needed or if the liquid vol i 40 m3; OK for high pressure. Cocurrent: surface area 50–400 m2/m3; Downflow: surface area 20–1000 m2/m3. Ha I I 0.3 and d+ = 4000–10 000. Can handle solids. Incurs a high pressure drop.

Gas–liquid–catalytic solid: Surface area 50–350 m2/m3.

Liquid–liquid: Surface area 7–75 m2/m3. Rotating disk contactor, RDC, can handle dirty fluids and large throughputs. Need flow ratios 1:1; difficulty in handling

6.13 PFTR: Bubble Reactor 237

systems with low interfacial tensions that tend to emulsify. Related topics solvent extraction, Section 4.10 and Size reduction, Section 8.3.

x Guidelines

Creation of bubbles. Bubbles can be created by injecting gas through holes in a pipe or sparger, injecting gas through a porous plate or diffuser, by introducing gas below an agitator by means of a sparger or a single injection tube or by inducing gas into a liquid via a jet or ejector. Diffuser Aeration: 0.3–0.5 dm3/s m3. 15– 30 dm3/s m2 diffuser area; or 1.5–6 dm3/s m of linear length of diffuser. Need 75–110 m3 air/kg BOD removed. 0.018–0.04 g oxygen absorbed/dm3 air sparged into the liquid. kL a = 0.0008 1/s. kLa = 0.02–0.08 1/s for power input 0.15–1 kW/ m3. Efficiency of oxygen transfer is, in general, 5–15 % with 8 % for porous tube diffusers and 6 % for coarse-bubble diffusers.

Typical diameters of bubbles formed by these different methods are given in Fig. 1.2. The usual range is 200 mm to 3 mm. In CSTR reactors the bubble diameter is usually 2 to 2.5 mm; for froth flotation of minerals, 1 mm; for foam fractionation the diameter is in the range 0.8–1 mm; for DAF, the bubbles nucleate on particulates, their diameter is usually 70–90 mm.

Gas–liquid: Can operate countercurrently or cocurrently. Holdup: volumetric liquid holdup per total reactor volume: i 0.7 and usually 0.95, Gas holdup = 0.05–0.4 increasing with increase in gas velocity. Superficial gas velocity 1–30 cm/s although it has been as high as 50 cm/s. Mass transfer coefficients: typical liquid mass transfer coefficient = kLa = 0.005–0.01 1/s; kL = 0.6– 0.7 q 10–4 m/s. For gas phase kG a = 1–3 1/s.

Surface area gas–liquid per volume of reactor: 20–1000 m2/m3 volume reactor depending on flow conditions.

Surface area gas–liquid per volume of liquid phase: 120–700 m2/m3 liquid phase volumetric ratio liquid to mass transfer liquid film, d+ i 100. Power: cocurrent: 0.03–2 kW/m3; countercurrent: 0.04–1 kW/m3. See Sections 1.6.1 or 1.6.3.

Bubble columns including bio: superficial gas velocity, 0.03 to 0.04 m/s; holdup I 0.2, kL a = 0.005 to 0.01 1/s. kL a is independent of the diameter if column diameter i 0.15 m; kL a is not affected by the type of gas sparger if the gas velocity exiting the orifice i 0.03 m/s. If I 0.03 m/s, then use a sintered plate. Height: 3 I height I 12 m; allow 0.75 of diameter or 1 m at the top for foam disengagement. Energy 0.01 to 3.5 kW/m3. OTR: 3.3 g/s m3; mixing time: 200 s; gas content: 30 %; maximum volume 5000 m3. We should note that the kLa values in practice differ from those reported for ideal systems such as air–water; air–nutri- ent; air–actual microbial system. These columns have been used for the production of steroids, acetic acid, SCP, single cell protein, beer, vinegar, yeasts, bacteria, mold fungi, baker’s yeast and for waste water treatment.

Air lift loop including bio: These are used often. They provide good mixing, efficient OTR (but not as high as for STR described in Section 6.27), low shear, excellent heat transfer, 100 000 to 500 000 L. kLa = 0.01–0.1 1/s for power input 0.5–1 kW/m3, kLa = 0.01–0.025 1/s for power input 0.2–0. 5 kW/m3. Working

238 6 Reactors

volume 2/3 total volume. H/D = 5 to 10. Power: 0.2–3.5 kW/m3; OTR: 2.7 g/s m3; mixing time: 80 s; gas content: 30 %; maximum volume 3000 m3. They have been used for culture broths of low viscosity, and for SCP: 30–40 kW/m3; remove heat with DT I 10 hC and keep reaction temperature about 37 hC.

Jet loop, with H/D = 5 to 20; jet velocity 20 m/s; internal draught tube. Power: 5 kW/m3; OTR: 2.8 g/s m3; mixing time: 60 s; gas content: 30 %; maximum volume 500 m3. This has been used for the mass cultivation of yeast on paraffin and the cultivation of bacteria.

External downcomer loop has been used by ICI pressure process for SCP. Deep shaft, tower loop, has been used for SCP from n-paraffins and for waste water treatment;

Gas–liquid–catalytic solid: including bioreactors. Catalyst diameter, I 0.1 mm. Operate semibatch. Holdup: volume fraction liquid 0.8–0.9; volume fraction catalyst 0.01; volume fraction gas 0.1–0.2. Gas holdup slightly less than for GL systems. Backmixing: solids backmixing Pe = 2–5 for superficial gas velocities of 2–7 cm/s; liquid phase backmixing about the same as GL systems; gas phase backmixing, about the same as GL systems. Surface area: surface area solid 500 m2/m3; surface area gas–liquid 100–400 m2/m3; power 0.1–2 kW/m3 sufficient to keep catalyst in suspension. Heat transfer solid wall to mixture i 1 kW/ m2 K; presence of solids increases the heat transfer coefficient. Catalyst activity: variable but often able to avoid diffusion limitations because of small diameter catalyst. Catalyst selectivity: OK. Catalyst stability: change catalyst between batches. Heat exchange OK. Consider complications because of catalyst deposition and erosion.

Liquid–liquid: creation of drops, see Sections 5.3 and 8.3. Superficial dispersed drop velocity 0.001–0.02 m/s with usual values 5.5 L/s m2. This bubble column gives the smallest reactor volume compared with STR. For example, for esterification: reactor volume divided by the daily production, m3 day/kg. PFTR 0.7 m3 day/kg. 3 CSTR in series, Section 6.26, 0.85 m3 day/kg. Batch STR, Section 6.27, 1.04 m3 day/kg. CSTR, Section 6.29, 1.22 m3 day/kg.

RDC: sum of the superficial velocities for both phases is 1–2.5 cm/s; diameter I 9 m for SX but usually I 2.5 m for reactions. Related topic SX, Section 4.10. Aeration basin: circular or rectangular basin of c/s or concrete with submerged air diffusers.

Liquid–solid with air used for mixing: example Pachuca leacher/reactor. L:D = 2.5:1.

x Good Practice

Electrolytes in the liquid alter the bubble diameter, the holdup, the interfacial area per unit volume in mechanically agitated devices and affects the kLa for bubble columns.

For gas–liquid, ensure operation is in the correct bubbling flow regime. See illustrative guide in Fig. 2.2, Section 2.4. Prevent foaming.

6.14 PFTR: Spray Reactor and Jet Nozzle Reactor 239

x Trouble Shooting

“Carryover”: [ foaming]*.

[Foaming]*: surfactants present/dirt and corrosion solids/natural occurring surfactants/pH far from the zpc/naturally-occurring polymers/insufficient disengaging space above the liquid/antifoam ineffective (wrong type or incorrect rate of addition)/bubble rate too high/mechanical foam breaker not rotating/baffle foam breaker incorrectly designed or damaged/asphaltenes present/liquid downflow velocity through the foam is too low. See Section 1.12 for generic causes of [ foaming]*.

See Trouble Shooting: STR, Section 6.27, for more on trouble shooting aerobic bioreactors.

6.14

PFTR: Spray Reactor and Jet Nozzle Reactor

Many options are available besides the traditional liquid gravity spray into a vessel filled with gas. These include options where the liquid is pumped through a jet or spray nozzle. Venturi jet the liquid and gas are mixed in the venturi. Plunging jet (buss loop reactor): vertical column filled about 2/3 with liquid, no internals. Liquid is withdrawn at the bottom and liquid is pumped into a top-mounted jet that induces air and resulting mix impinges vertically under the top surface of the liquid. (This is in contrast with the jet loop, Section 6.13, where the ejector is at the bottom of the liquid column.)

The external circulating nozzle does the GL mixing outside the column and injects the mixture into the top of the vessel. This is a vertical column with no internals. Liquid is withdrawn at the bottom and pumped through a venturi that induces air. The air–liquid mixture enters at the top of the column.

The internal circulating nozzle, or immersed column, is similar to the external circulating nozzle except that all the action takes place within the column. This is a vertical column with a small diameter central axial injection tube. Liquid is pumped from a submerged pump to the top of the tube; compressed air is introduced at the top of the tube and the two phase mixture is injected down the tube and into the bottom of the vessel.

x Area of Application

Phases: GL, LL, GLS (bio)

Gas liquid and GLS (bio): Residence time, very short. Reaction rates, very fast; need rapid absorption. Reaction is controlled by mass transfer. Very high gas capacity. Used for neutralization reactions with one of the reactants in the gas phase. Ha i 3 and d+ = 2–10. For surface area see Sections 1.6.1 and 1.6.3.

Gravity spray: surface area 30–70 m2/m3; target species Henry’s law constant 103 to 104 kPa/mol fraction; feed gas concentration 0.3 to 4 vol %. Can handle foaming and solids-laden gases, low pressure drop. Reactions: good for reaction with highly soluble gases.

240 6 Reactors

Venturi jet nozzle: Surface area: 200–2500 m2/m3; very soluble gas only with target species Henry’s law constant I 103 kPa/mol fraction; feed gas concentration i 1 vol %.

Liquid–liquid: Surface area 7–75 m2/m3. Related topics: solvent extraction, Section 4.10 and size reduction, Section 8.3.

x Guidelines

Creation of the gas–liquid contact surface via sprays or venturi is discussed in Section 8.2 with surface area and power input information given in Sections 1.6.1 and 1.6.3. In general, for sprays created by a venturi the drop size is 1–30 mm. Gravity Spray towers: superficial velocity about 5.5 L/s m2; mass transfer coefficients liquid phase: 1.5 q 10 –4 I kL I 3 q 10 –4 m/s; for the gas phase 0.4 I kG I 1 mol/m2 atmos s or in other units 0.01 I kGRT I 0.25 m/s; kLa = 0.0007 to 0.015 1/s. critical energy consuming phase is the liquid atomization; gas energy 8 kJ/m3; with liquid to gas ratio high; design on gas phase controlling. Superficial gas velocity 0.75–2 m/s and usually 1 m/s, holdup I 0.8. Power usage 0.03 to 0.5 kW s/m3. Dp gas = 0.6 to 1.2 kPa. Related topic gas–solid separation, Section 5.2 and the use of spray nozzles to create a dispersion, Section 8.2.

Gas–liquid: Venturi: Dp gas = 1–6 kPa and usually 5 kPa, velocity in the throat 30–100 m/s and usually 100 m/s; mass transfer coefficient for liquid phase kL = 7 q 10 –4 m/s; for gas phase kG = 10–2–3 q 10–2 m/s; gas–liquid surface area 150–300 m2/m3. For fluids with surface tension 40–70 mN/m and viscosity of liquid = 1 mPa s; critical energy consuming phase is the gas at about 20 kJ/m3 with liquid to gas ratio about 1.3 to 1.6 L/m3; design on gas phase controlling. Power usage 0.04–8 kW s/m3. Related topic Gas–Solid Separation, Section 5.2 and Size Reduction, Section 8.2.

For gas–liquid–solid:

Plunging jet: 8 to 12 m/s. Power requirement high but is reduced by increasing the volume of reactor: Power: 3–8 kW/m3; OTR: 3.3 g/s m3 (kLa up to 0.3 s–1) ; gas content: 20 %; maximum volume 200 m3.

External circulating nozzle: Power: 6 kW/m3; OTR: 0.8 g/s m3; mixing time: 50 s; gas content: 30 %; maximum volume 200 m3.

Internal immersed column: Power: 0.08 kW/m3; OTR: 0.003 g/s m3; gas content: 2 %; maximum volume 4000 m3.

Liquid–liquid: Superficial velocity = 50 % flooding. See also Section 1.6.2.

6.15

PFTR: Trays

The general characteristics of gas–liquid contacting are described in Section 1.6.1. Other operations that use this type of contactor include gas absorption, Section 4.8, gas desorption/stripping, Section 4.9; gas–liquid separations, Section 5.1; turbulent bed contactor (TCA, TVA) contactor, Section 5.2, distillation, Section 4.2, reactive distillation, Section 6.35 and direct contact heat exchange Sections 3.8 and 3.9.